Process for the recovery of nickel and / or cobalt from a concentrate

ABSTRACT

A process for the recovery of nickel and/or cobalt values from a concentrate containing nickel or cobalt hydroxide comprises the steps of subjecting the concentrate to a leaching stage with an ammonium solution to produce a leach solution containing nickel and/or cobalt values and a residue. The nickel concentration in the leach solution is controlled to a maximum value of about 3 to 25 g/l, preferably 8 to 15 g/l and more preferably 10 g/l.

CROSS REFERENCE TO RELATED APPLICATION

[0001] This application is a division of U.S. patent application Ser.No. 08/660,290 filed Jun. 7, 1996, which is a continuation-in-part ofU.S. patent application Ser. No. 08/488,128 filed Jun. 7, 1995 which, inturn, is a continuation-in-part of U.S. patent application Ser. No.08/425,117 filed Apr. 21, 1995 which, in turn, is a continuation-in-partof U.S. patent application Ser. No. 08/098,874 filed Jul. 29, 1993 whichissued as U.S. Pat. No. 5,431,788 on Jul. 11, 1995. The contents of theforegoing applications are incorporated herein by reference.

FIELD OF THE INVENTION

[0002] This invention relates to a process for the recovery of nickeland/or cobalt from a concentrate.

BACKGROUND OF THE INVENTION

[0003] Nickel sulphide ores are presently treated in commercial practiseby a variety of processes in which the first step is almost always aphysical concentration by flotation to upgrade the Ni content,typically, from a range of 0.5% to 2.0% up to 7 to 25% Ni, as aconcentrate. The subsequent treatment of this concentrate is usuallypyrometallurgical (smelting) to produce a Ni matte or an artificial highgrade sulphide with about 20% to 75% Ni.

[0004] The matte is then generally refined to nickel products byhydrometallurgical techniques.

[0005] This combination of pyrometallurgical/hydrometallurgicalprocessing of Ni concentrates is now well established commercially witha number of variations, particularly in the hydrometallurgical portion.Most processes recover some portion of the associated metal values wherepresent, such as copper and cobalt. In addition, a leach residuecontaining precious metals, such as gold and silver, as well as platinumgroup elements, e.g. platinum and palladium, is often produced forsubsequent recovery of contained values.

[0006] This treatment scheme has some inherent drawbacks. Thoseassociated with the pyrometallurgical step, include:

[0007] (i) Production of smelter gases including SO₂, which must now betreated in an acid plant to produce sulphuric acid byproduct, whichfrequently is difficult to market from a remote location. (The capitaland operating costs of such acid plants impact on the overall economiesof the process.)

[0008] (ii) Losses of nickel and particularly cobalt into the slagproduced during smelting, often more than 50% of cobalt input.

[0009] (iii) High costs of smelting in general, particularly for lowgrade concentrates (<10% Ni).

[0010] (iv) Difficulty in treating certain concentrates with deleteriouselements, such as magnesium (Mg) and arsenic (As).

[0011] The hydrometallurgical steps for treating Ni matte varyconsiderably but all known commercial processes have one or more of thefollowing disadvantages:

[0012] (i) High costs for reagents such as caustic soda or ammonia,required for neutralization.

[0013] (ii) Large byproduct production, such as ammonium sulphate orsodium sulphate, which are difficult to market.

[0014] (iii) High energy costs, due to large temperature changes duringthe process.

[0015] (iv) Complex and costly process flowsheet, leading to highcapital and operating costs.

[0016] As an alternative to the establishedpyrometallurgical/hydrometallurgical route outlined above, there is oneknown process using wholly hydrometallurgical steps, that treatsconcentrates without smelting. It uses a pressure leaching techniquewith ammoniacal solution. This avoids most of the disadvantagesassociated with the smelting processes, but unfortunately still suffersfrom all of the listed disadvantages of the known hydrometallurgicalroutes, and in fact is not even as efficient overall as the best of thepyrometallurgical/hydrometallurgical routes.

[0017] Copper or nickel sulphide ores often also contain other metalvalues, such as cobalt, as well as precious metals, such as gold andsilver and the platinum group metals. Since these ores are typically lowgrade ores, in so far as copper/nickel is concerned, and also have ahigh sulphur to copper/nickel ratio, the economical extraction ofcopper, nickel and cobalt values have been problematical. Some sulphideores contain such low copper/nickel values that the recovery of preciousmetals must be high in order to render the process economical. Due tothe pyrite content of some ores, the recovery of gold by conventionalcyanidation is often difficult, which also renders the treatment of theore uneconomical.

SUMMARY OF THE INVENTION

[0018] According to the invention there is provided a process for therecovery of nickel and/or cobalt values from a concentrate containingnickel and/or cobalt hydroxide, comprising the steps of subjecting theconcentrate to a leaching stage with an ammonium solution to produce aleach solution containing nickel and/or cobalt values and a residue; andcontrolling the concentration of nickel in the leach solution to amaximum value of about 3 to 25 g/l.

[0019] The term “concentrate” in this specification refers to anymaterial in which the metal value content has been increased to a higherpercentage by weight as compared with the naturally occurring ore andincludes man made artificial sulphide ore, such as matte, and metalvalues precipitated as solids such as hydroxides and sulphides.

[0020] Further objects and advantages of the invention will becomeapparent from the description of preferred embodiments of the inventionbelow.

BRIEF DESCRIPTION OF THE DRAWINGS

[0021]FIG. 1 is a flow diagram of a hydrometallurgical metal extractionprocess.

[0022]FIG. 2 is a flow diagram giving more details about the solventextraction steps of the process of FIG. 1.

[0023]FIGS. 3A and B show a flow diagram of a further embodiment of theprocess for the recovery of precious metals.

[0024]FIG. 4 is a flow diagram of another hydrometallurgical metalextraction process.

DETAILED DESCRIPTION OF PREFERRED EMBODIMENTS

[0025] The process is suitable for the treatment of copper ores,particularly copper sulphide ores, which also contain nickel and/orcobalt values, or nickel/cobalt sulphide ores without significant coppervalues, as well as nickel/cobalt oxide (laterite) ores. In addition, theprocess can treat nickel/cobalt ores with other elements oftenconsidered to be deleterious, such as magnesium, arsenic and zinc, orelements which are valuable and worth recovery, such as the preciousmetals, gold and silver, and the platinum group metals.

[0026] The feed ore or concentrate to the process may contain one ormore sulphide minerals of the base metals Cu, Ni, Co and Zn, frequentlycombined with Fe and sometimes with other elements such as As, Sb, Ag,etc.

[0027] Typical sulphide minerals of the base metals listed above are:Copper: Cu₂S - Chalcocite, CuFeS₂ - Chalcopyrite Nickel: NiS -Millerite, (Ni,Fe)₉S₈ - Pentlandite Cobalt: Co₃S₄ - Linnaeite,(Co,Fe)AsS - Cobaltite Zinc: ZnS - Sphalerite, (Zn,Fe)S - Marmatite

[0028] The metal:sulphur ratio in this context is the ratio of the totalbase metals (Cu, Ni, Co, Zn) to sulphur in the concentrate, and this isa measure of the grade of the concentrate.

[0029] Typically the metal:sulphur ratio varies from 1.5 for high gradeconcentrates down to 0.2 for low grade concentrates. For concentratesthat are predominantly Ni/Co, the metal:sulphur ratio is more often inthe lower part of the range, from 0.2 to 0.8 (Fe is specificallyexcluded from this calculation, even though it is present in practicallyall sulphide concentrates).

[0030] The significance of the metal:sulphur ratio to the process, isthat if affects the metallurgy occurring during the initial operation ofpressure oxidation.

[0031] The different embodiments of the process may be used to treat arange of Ni/Co concentrates in which the metal:sulphur ratio varies fromlow to high as outlined above. However, in addition to this ratio, thereis another important characteristic which must be taken into account.The degree of sulphur oxidation (to sulphate) during pressure oxidation.Sulphur contained in concentrate is converted during pressure oxidationeither to elemental sulphur (S°) (no sulphur oxidation), or oxidized tosulphate (SO₄ ⁼). Typically about 70-95% of the sulphur is not oxidized,and is produced as elemental sulphur. Expressed another way, sulphuroxidation (to sulphate) varies usually from 5 to 30%. It is consideredbeneficial to minimize sulphur oxidation, and it is an importantobjective of this process to do so. This is facilitated by theintroduction of a source of sulphate or bisulphate, such as H₂SO₄, intothe pressure oxidation stage.

[0032] The significance of sulphur oxidation is that it produces acid,which must eventually be neutralized, and it affects the distribution ofCu, Fe and other elements in the product slurry from pressure oxidation.Higher acid slurries (low pH) contain Cu in solution, whereas lower acidslurries (high pH) have Cu in solid form, as basic copper sulphate.

[0033] For concentrates with low metal:sulphur ratio and/or high sulphuroxidation, the process flowsheet shown in FIG. 1 is the general case.This is referred to as Mode C. Enough acid is produced during pressureoxidation 12, that it is necessary to neutralize this acid by slakedlime in the latter stages of the autoclave. This is indicated as theneutralization 501 in FIG. 1. Without this neutralization, the productslurry would have low pH, resulting in significant Fe in solution, andalmost all of the Cu as well.

[0034] It is an important feature of the process that this productslurry contain minimal Fe in solution (less than 100 ppm) and about 1-5g/l Cu in solution. By adjusting the amount of slaked lime added in theneutralization 501, these objectives can be achieved even withconcentrates that have low metal:sulphur ratio and exhibit relativelyhigh sulphur oxidation, e.g. 15-30%. A typical example of this type ofconcentrate is a pentlandite/pyrite type of mineral assemblage.

[0035] However, for concentrates that have high metal:sulphur ratioand/or low sulphur oxidation, the total amount of acid produced duringpressure oxidation 12 is less, and no neutralization 501 may be requiredto achieve a product slurry with low Fe and Cu in the desired range.This embodiment of the process is termed Mode A and is described belowwith reference to FIG. 4. A typical example of this type of concentrate,is a pentlandite/chalcopyrite/pyrrhotite type of mineral assemblage.

[0036] In Mode A, the amount of acid consumed during pressure oxidationby other chemical reactions is more than sufficient to use up all theacid produced by sulphur oxidation.

[0037] Examples of both Mode A and Mode C required for two differentconcentrates are shown in the table below: CONCENTRATE PROCESS ASSAY %METAL:SULPHUR % SULPHUR TYPE Cu Ni Co S RATIO OXIDATION MODE A 6.3 140.6 34 0.61  6 MODE C 0.1 22 0.6 29 0.78 15

[0038] Thus, the first concentrate with 14% Ni exhibited only 6% Soxidation in pressure oxidation, and thus was treated by Mode A, whereasthe second concentrate required Mode C, due to the higher S oxidation(15%).

[0039] Process Mode C will now be described with reference to FIG. 1.

[0040] First the ore or concentrate is subjected to pressure oxidation12 in an autoclave in the presence of an acidic solution containingsulphate, chloride and copper ions. In the present example the amount ofH₂SO₄ introduced into the autoclave is about 40 g/l and theconcentration of chloride in solution is about 10-12 g/l. Typically thetemperature is about 90° C. to about 160° C. under an oxygen partialpressure of about 200-2000 kPa. The retention time is about 0.5-5.0hours, depending inversely on temperature, and the process is normallycarried out in a continuous fashion in the autoclave. However, theprocess can also be carried out in a batch-wise fashion, if desired.

[0041] The neutralization 501 is effected by pumping slaked lime intothe last one or two compartments at the exit side of the autoclave, atabout 10-20% solids in water.

[0042] After pressure oxidation 12, the slurry produced in the autoclaveis discharged through one or more flash tanks 22 to reduce the pressureto atmospheric pressure and the temperature to 90-100° C.

[0043] The slurry is then further cooled and subjected to filtration 24to produce a pressure oxidation filtrate 29 and a solid residue(pressure oxidation filter cake).

[0044] The neutralization step 501 is used to precipitate soluble copperinto the pressure oxidation filter cake, that would otherwise report tothe pressure oxidation filtrate 29. Thus, the neutralization 501 can beused to minimize copper in the filtrate 29, typically down to 1 to 5 g/lcopper, which makes the subsequent removal of copper from solutioneasier. In addition, the neutralization 501 helps to minimize Fe in thepressure oxidation filtrate 29. However, when adding slaked lime it ispreferable not to add too much so as to precipitate Ni/Co. Typically,adding slaked lime so that the pressure oxidation filtrate 29 has a pHof between about 3 and 4 has been found suitable for removing most ofthe copper and yet minimizing Ni/Co precipitation.

[0045] The pressure oxidation filtrate 29 is generally subjected tocopper solvent extraction 50, particularly if significant copper valuesare present in the original concentrate, to recover the copper valuesand to reduce [Cu²⁺] in the raffinate 63 as low as possible, typicallyless than 100 ppm. In addition, the pressure oxidation filter cake issubjected to an atmospheric leach 14 to recover copper in solution,which solution is subjected to Cu solvent extraction 16. The leach 14 iscarried out with raffinate 120 from the Cu solvent extraction 16 whichis dilute acid at about 3-20 g/l H₂SO₄. In addition the leach 14 helpswash the entrained solution containing any Ni/Co values out of thepressure oxidation filter cake. These values which will accumulate instream 51 can be recovered on a bleed basis (say 1 to 10% of flow,depending on concentration) by precipitating at pH 7 to 8 with slakedlime as Ni/Co hydroxides, similar to the conditions in the precipitation506, described below. The mixed Ni/Co hydroxide can then be filtered offand recycled to a purification stage 500, described below.

[0046] The slurry 31 resulting from the leach 14 is difficult to filterand liquid/solid separation is effected by means of a series ofthickeners in a counter current decantation (CCD) arrangement 34. Washwater is provided by a portion of the raffinate from the solventextraction 16, which is split at 36 and neutralized at 46 usinglimestone to remove acid. The slurry from the neutralization 46 isfiltered at 48, to produce a gypsum residue and the liquid 51 isrecycled as wash water.

[0047] The loaded extractant from the solvent extractions 50 and 16 issubjected to stripping 44 and is then sent to copper electrowinning 20.

[0048] The Cu solvent extractions 50 and 16 are operated with a commonextractant. This is shown in FIG. 2 where the broken line indicates theorganic extractant being circulated after stripping 44. The stripping 44is effected with spent acid or electrolyte 55 from the electrowinning 20to obtain a pure copper sulphate solution or pregnant electrolyte 57which is then passed to the electrowinning stage 20. Any suitable copperextractant capable of selectively removing Cu from an acid solution alsocontaining Ni/Co/Zn/Fe, may be used. An extractant that is found to besuitable is a hydroxy-oxime, such as LIX 84™ or LIX 864™ reagents fromHenkel Corporation.

[0049] If no significant copper values are present in the ore orconcentrate, it is nevertheless beneficial to carry out the pressureoxidation 12 in the presence of copper ions (e.g. 5 to 10 g/l Cu).Copper ions can be added in the form of a copper salt, such as CuSO₄ orCuCl₂. Thereafter, Cu solvent extraction and stripping are still carriedout but the electrowinning 20 will be omitted and the pregnant copperliquor resulting from stripping 44 of the organic extractant will berecycled to the pressure oxidation 12. Alternatively, a copperconcentrate can be added in which case the copper can be recycled afterCu solvent extraction and stripping or sent to electrowinning forrecovery of the copper. This will also be the case if a laterite ore isbeing processed.

[0050] The raffinate 63 is subjected to a purification stage 500, toprepare a solution of Ni/Co free from elements such as Fe, Zn and Cuthat cause difficulty in the subsequent process steps of solventextraction and electrowinning of Ni and Co. The purification stage 500is a precipitation step in which residual Cu, Fe and Zn are precipitatedby the addition of slaked lime and recycled Mg(OH)₂. Typically, the feedsolution to the purification stage 500 will contain copper and iron, aswell as any zinc and magnesium present in the concentrate. Theprecipitation 500 is effected at a pH of about 5 to 6 so that, ideally,no more than about 1 ppm Zn, 1 ppm Cu and 1 ppm Fe remain in thesolution. It is also important not to precipitate too much Ni/Co. Thisis achieved by careful control of pH, i.e. not allowing the pH to risetoo high. The recycled Mg(OH)₂ has been found to be beneficial in thisregard.

[0051] The product from the precipitation 500 is subjected to aliquid/solid separation 502. The Cu, Fe and Zn, which precipitate ashydroxides, can be reprocessed by a dilute acid wash or leach 503,particularly for Ni/Co recovery. The product from the acid wash 503 issubjected to a liquid/solid separation 505 leaving principally Cu, Feand Zn hydroxides, which provides an outlet for zinc from the system.The liquid 504 from the liquid/solid separation 505, is recycled to thepressure oxidation 12.

[0052] If the Zn content is sufficiently high, the Cu/Fe/Zn hydroxidecan be further leached with dilute acid to selectively recover zinc. Inan extreme case, a zinc solvent extraction step can be included, ifdesired.

[0053] The concentrations of Ni, Co and Mg in solution after theprecipitation 500 will depend on the composition of the concentrate.Depending on the mineralogy, it is possible that most of the magnesiumin the concentrate leaches during the pressure oxidation 12. Thus, forNi/Co concentrate containing say 20% nickel and 5% magnesium, thetypical solution after the precipitation 500 will be about 30 g/l nickeland about 6 g/l magnesium. The magnesium content will be greater in thecase of a laterite ore.

[0054] The solution resulting from the liquid/solid separation 502, issubjected to a selective precipitation step 506 in which Ni and Co areprecipitated as hydroxides or carbonates with a suitable neutralizationagent, such as slaked lime (Ca(OH)₂), soda ash (Na₂CO₃), ammonia orcaustic soda (NaOH). This is effected at a pH of about 7 to 8, whilstminimizing the precipitation of Mg(OH)₂. A preferred neutralizationagent is slaked lime due to its relatively low cost, and because thereaction does not introduce any new cations, such as Na⁺ and NH₄ ⁺, intothe liquor.

Neutralization with Slaked Lime

[0055] $\begin{matrix}\left. {{{NiSO}_{4}({aq})} + {{Ca}({OH})}_{2}}\rightarrow{{{{Ni}({OH})}_{2}(s)} + \underset{({gypsum})}{{{CaSO}_{4} \cdot 2}H_{2}{O(s)}}} \right. & (1)\end{matrix}$

[0056] A similar reaction occurs with CoSO₄ and MgSO₄, producing Co(OH)₂and Mg(OH)₂ respectively.

Neutralization with Caustic Soda) (NaOH)

NiSO₄(aq)+NaOH→Ni(OH)₂(S)+NaSO₄(aq)  (2)

[0057] However, it is important to have some Mg present in theprecipitated solid, which facilitates the separation of Ni and Co, aswill be described below. A two-stage counter current precipitationsequence has been found beneficial.

[0058] In some circumstances, a precipitation with caustic soda orammonia for instance that does not produce a solid byproduct (gypsum) isadvantageous, so that the Ni precipitate is of a higher grade, and freefrom calcium.

[0059] The product from the precipitation step 506 is subjected to aliquid/solid separation 508.

[0060] The liquid from the liquid/solid separation 508 is subjected to aprecipitation step 510, preferably again with slaked lime, for the samereasons as above, to precipitate additional Mg, if needed, thereby toprevent accumulation of Mg in the system. The product from theprecipitation step 510 is subjected to a liquid/solid separation 512.The solid from the separation 512 is a magnesium hydroxide byproduct514. As indicated above, some of the magnesium hydroxide byproduct 514is recycled for use in the precipitation 500. The liquid from theseparation 512 is recycled to the pressure oxidation 12, as indicated bythe recycle stream 516.

[0061] The solid hydroxide cake from the separation step 508, containingthe Ni and Co values, is subjected to a leach 518 with an ammoniumsolution at a pH of about 6 to 8.

[0062] The ammonium solution may be ammonium sulphate or ammoniumcarbonate but the former has been found to be superior because it has alower pH, thus allowing for a better Co to Ni separation in solution. Inaddition, ammonium sulphate has a lower ammonia (gas) vapour pressure,and as well, the Ni/Co extractions are superior with ammonium sulphate.In the present example a 200 g/l ammonium sulphate solution is used.

[0063] The reactions which take place during the leach 518, in whichsoluble nickel and cobalt diammine sulphates are formed, are as follows:

(NH₄)₂SO₄+Ni(OH)₂→Ni(NH₃)₂SO₄+2H₂O  (3)

(NH₄)₂SO₄+Co(OH)₂→Co(NH₃)₂SO₄+2H₂O  (4)

[0064] The Mg present in the solid also dissolves, as follows:

(NH₄)₂SO₄+Mg(OH)₂→MgSO₄→2H₂O+2NH₃  (5)

[0065] In carrying out the leach 518, it is not attempted to leach out100% of the Ni/Co values in the solid but only about 90-99%. Thisenables the leach 518 to be carried out at a low pH rather than a higherpH of about 9 which would otherwise be required. This higher pH requiresthe addition of ammonia to the leach as a second reagent with theammonium sulphate.

[0066] A further problem which arises is that the known or commerciallyavailable Co extractant does not function effectively at this high pHvalue. The extractant degrades and it is not selective against Ni. As aresult, it is necessary to effect Ni extraction first, rather than Coextraction, which would then require reducing the pH by the addition ofa further reagent such as acid, which would in turn mean production ofbyproduct ammonium sulphate and consumption of the reagent ammonia.Another problem that arises is that, in order to effect Ni solventextraction first, it is necessary first to oxidize all the Co to the +3oxidation state to avoid extraction of Co with Ni. This oxidation isdifficult to achieve quantitatively. This, therefore, results in furtherprocess complications. Also it is necessary to reduce the Co³⁺ back toCo²⁺ following Ni extraction and this is equally difficult to achieve.

[0067] To avoid the above difficulties, the process according to thepresent invention provides effecting the leach 518 at a pH of about 6 toabout 8 and then subjecting the resultant solid to a subsequent washingstage 520 with dilute ammonium sulphate solution, as will be describedbelow.

[0068] A further aspect of the process is that the concentration ofnickel ions in solution during the leach 518 is controlled to remain ata relatively low value of about 10 g/l maximum. It has been found thatthis results in better Ni recovery during the leach 518. With the amountof Ni present in the solid known, the appropriate volume of liquidrequired to arrive at the desired Ni concentration can be calculated.

[0069] The product from the leach 518 is subjected to liquid/solidseparation 522.

[0070] The liquid from the separation 522 is subjected to a Co solventextraction 534 to provide a Co loaded extractant and a raffinate whichis then subjected to a Mg solvent extraction 536 to provide a Mg loadedextractant and a raffinate which is subjected to a Ni solvent extraction538 to provide a Ni loaded extractant and a raffinate.

[0071] The raffinate from the Ni solvent extraction 538 is recycled tothe leach 518.

[0072] The solid product from the liquid/solid separation 522 issubjected to the repulp or washing step 520 as indicated above where thesolid is washed with ammonium sulphate solution. This is a weak ammoniumsulphate solution of about 10% the concentration of the solution of theleach 518. It results from the washing of entrained ammonium sulphatesolution from the solid in the washing step 520.

[0073] The product from the repulp step 520 is subjected to aliquid/solid separation 524 and the solid is washed with water. The washwater and liquid from the liquid/solid separation 524 is subjected to aCo solvent extraction 526 to again provide a Co loaded extractant and araffinate which is subjected to Mg solvent extraction 527 to provide aMg loaded extractant and a raffinate which is subjected to a Ni solventextraction 528 to provide a Ni loaded extractant and a final raffinatewhich is recycled to the repulp step 520.

[0074] To compensate for the water added during the water wash at theseparation 524, there is a bleed of the final raffinate to the strongammonium sulphate raffinate coming from the Ni solvent extraction 538.For this purpose, the strong ammonium sulphate circuit includes anevaporation step 539 to compensate for the raffinate bleed from the weakammonium sulphate raffinate.

[0075] The Co solvent extractions 534, 526, the Mg solvent extractions536, 527 and the Ni solvent extractions 538, 528, respectively, are alloperated with a common extractant, as is the case with the Cu solventextractions 50, 16.

[0076] An extractant which has been found to be suitable for both Co andMg extraction is an organic phosphorous acid extractant, morespecifically an organic phosphinic acid based extractant, such as Cyanex272™, of Cyanamid Inc., which comprises bis 2,4,4-trimethylpentylphosphinic acid. For the Ni extraction, a hydroxy-oxime basedextractant, such as LIX 84™, of by Henkel Corp, has been found to besuitable.

[0077] The respective Co, Ni and Mg loaded extractants are scrubbed withsuitable aqueous solutions to remove entrained ammonium sulphatesolution and then stripped with dilute acid to produce pure pregnantsolutions of Co and Ni and a Mg pregnant liquor containing small amountsof Co and Ni. The Co and Ni solutions are sent to the Co and Nielectrowinning stages 530 and 532, respectively.

[0078] Prior to stripping, the Co loaded extractant is scrubbed with aCo concentrate solution which is split off from the Co pregnant solutiongoing to Co electrowinning and/or a Mg concentrate solution which issplit from the Mg pregnant liquor. This is to facilitate the removal ofNi values which may be present in the Co loaded extractant. Likewise,the Mg loaded extractant can be scrubbed with a Mg concentrate solutionwhich is split off from the Mg pregnant liquor.

[0079] For good separation of Co from Ni during Co solvent extractionand Ni solvent extraction, it has been found beneficial to have some Mgpresent in the solution feed to the Co solvent extraction. Typically,solution analysis has the same ratio of Co to Ni as found in theoriginal feed concentrate (commonly 1:30). Thus for 10 g/l Ni, 0.33 g/lCo is typical.

[0080] The same extractant is used for both the Co and Mg solventextractions 534 and 536. The extractant is more selective for Co thanfor Mg, and more selective for Mg than for Ni. During the Co solventextraction 534, the amount of extractant used is limited to occupy allthe available sites with Co ions, to a major extent, and with Mg ions,to a lesser extent, which counteracts the extraction of Ni. During theMg solvent extraction 536, the available sites are filled with mainly Mgions and, to a lesser extent, with some Co ions and possibly also asmall amount of Ni ions. The Ni and Co ions are then recovered by therecycle of the Mg pregnant liquor to the Ni/Co precipitation 506, asindicated by the arrow 543.

[0081] It has further been found beneficial to maintain a Mgconcentration about equal to the Co concentration, although this mayvary quite widely from say 1:5 to 5:1.

[0082] The benefit of having Mg present is that:

[0083] (i) it minimizes the amount of Ni that is extracted during Cosolvent extraction, whilst allowing

[0084] (ii) high Co percent extraction, i.e., greater than 90%, and

[0085] (iii) a high Co to Ni ratio in the Co product, i.e., Co :Ni>1000:1.

[0086] Without Mg present, some compromise must be reached in the Cosolvent extraction, whereby

[0087] (i) some Ni is co-extracted with Co, or

[0088] (ii) the Co extraction is incomplete, or

[0089] (iii) the Co to Ni ratio in the Co product is too low.

[0090] With Mg present, some Co (i.e. 5-10%) can be left un-extractedduring Co solvent extraction and instead will be extracted during Mgsolvent extraction. The products of Mg solvent extraction are:

[0091] (a) Pregnant liquor from stripping containing some Mg, Ni and Co,which is recycled and not lost; and

[0092] (b) Mg raffinate with very low Co levels, i.e. about 1 ppm, whichallows the subsequent Ni solvent extraction to produce a very good Ni toCo ratio in the Ni pregnant liquor going to Ni electrowinning. Thus,very pure Ni cathodes and Co cathodes result.

[0093] The solid from the liquid/solid separation 524 is washed (540)with dilute acid to recover entrained Ni/Co which is recycled to theprecipitation 500. The solid residue after the liquid/solid separation542 is discarded.

[0094] A suitable temperature range for the Ni/Co leach 518 and Ni/Cosolvent extractions has been found to be about 30° C. to 60° C.,preferably about 40° C. to about 50° C.

[0095] Turning now to FIGS. 3A and B, the recovery of precious metals,such as gold and silver, will be described. This process involves thetreatment of the final residue stream 35 in FIG. 1.

[0096] The precious metals are not leached during the pressure oxidationstage 12 but remain in the solid residue 35 remaining after theatmospheric leaching stage 14.

[0097] In order to facilitate precious metal recovery, the flash down 22from the pressure oxidation stage 12 is carried out in two stages. Thefirst stage is at a temperature slightly above the freezing point ofelemental sulphur, i.e. about 120° to 130° C. with a corresponding steampressure of about 50-150 kPa. The process is preferably carried out in acontinuous mode, the retention time at the first flash let-down stagebeing about 10 to 30 minutes.

[0098] The second flash let-down stage is at atmospheric pressure andabout 90 to 100° C. with a retention time of again at least 10 minutes.This allows the elemental sulphur, which is still molten in the firstflash-down stage, to convert to one of the solid phases, such as thestable orthorombic crystalline phase. This procedure facilitates theproduction of clean crystals of elemental sulphur, which is important tothe recovery of the precious metals from the leach residue.

[0099] The leach residue 35 now produced by the atmospheric leachingstage 14 contains, in addition to the precious metals, hematite,crystalline elemental sulphur, unreacted sulphides (pyrite) and anyadditional products that may result from the particular concentratebeing used, e.g. gypsum and iron hydroxides.

[0100] Gold in the residue 35 is believed to be largely untouched by theprocess so far and most likely is in the native state. Silver, however,is oxidized in the pressure oxidation stage 12 and is probably presentas a silver salt, such as silver chloride or silver sulphate.

[0101] It has been found that conventional cyanidation does not leachgold well from the residue 35. It is believed that this is due to theencapsulation of the gold in mineral particles, such as pyrite. The goldcan however be liberated by the pressure oxidation of these minerals,referred to as “total oxidative leaching”. In order to effect suchleaching without oxidizing elemental sulphur also contained in theresidue 35, the process comprises the step of removing as much of theelemental sulphur as possible.

[0102] Firstly, by virtue of the two stage flash-down, good qualitysulphur crystals are produced. Secondly, the leach residue 35 issubjected to froth flotation 402 to produce a sulphur rich flotationconcentrate 404 and a sulphur depleted flotation tail 406. The tail 406is subjected to a solid/liquid separation 408 to produce a liquid whichis recirculated to a conditioning tank 410 upstream of the flotationstep 402 and a solid 412 which is sent to the total oxidative leachingstage 414.

[0103] The flotation concentrate 404 is filtered (416), and dried to alow moisture in a dryer 418. The product is then subjected to a sulphurleaching step 420 with a sulphur extractant. Any suitable sulphurextractant such as perchloroethylene (PCE) or kerosene may be used. Inthe present example hot PCE is used. The slurry from the leach 420 isfiltered 422 and the resulting liquid is subjected to cooling 424 toproduce crystalline S° and then filtered (425). The cooled sulphur canbe subjected to an optional sulphur purification step (not shown) toremove impurities, such as selenium and tellurium, therefrom. The solidsulphur is dried in a dryer 426 to produce a sulphur product 428. Theliquid from the filtration 425 is recycled to the hot PCE leach 420.

[0104] The solid residue from the filtration 422 is dried in a dryer430. The resulting product, which is a low sulphur residue 432, is sentto the total oxidative leach 414.

[0105] The PCE vapours from the cooling 424 and the dryers 426 and 430are recycled to the hot PCE leach 420 via a condenser 434.

[0106] A test was carried out in which 100 g of residue from theatmospheric leach 14 containing 25.1% elemental sulphur (S°) and 3%sulphide was processed through flotation 402 and leaching 420. Thisproduced 73.8 g of desulphurized residue (feed material for the totaloxidation leach 414) containing 1.9% S° and 4.1% sulphide, i.e. a totalof 6% total sulphur.

[0107] The desulphurized residue contained 5.9% of the elemental sulphur(S°) in the original leach residue, i.e. 94.1% was recovered to a pureelemental sulphur product.

[0108] The total oxidative leach 414 is carried out at about 200°C.-220° C. and 200-2000 kPa oxygen partial pressure, sufficient to fullyoxidize all sulphur and metal compounds to the highest valences,respectively. Thus all sulphur and pyrite are oxidized to sulphate. Theoxidation is conducted in acidic conditions, such as with the acid beingproduced in situ. If sufficient pyrite is present, the reaction ishighly exothermic and generally the desired operating temperature can beachieved. Typically about 10% of total oxidizable sulphur will besufficient with normal percentage solids in the feed slurry.

[0109] After the total oxidative leaching 414, the slurry is subjectedto neutralization 437 at pH 2-3 with limestone and then subjected to aliquid/solid separation 438 by means of a counter current decantation(CCD) circuit, to obtain a solid containing precious metals and a liquid13 which may contain base metal values, such as copper. The liquid 13can be combined with the liquid (stream 33) going to the solventextraction 16 for the recovery of copper, as indicated in FIG. 1.

[0110] A portion of the neutralized stream 51 (FIG. 1) of the raffinatefrom the Cu solvent extraction 16 is split off at 49 and the resultingstream 53 is partly used (about 80%) as wash water in the liquid/solidseparation 438 and partly recycled (about 20%) to the total oxidativeleach 414, as indicated in FIG. 3B. The precious metals recovery circuitof FIGS. 3A and B is indicated by the block 155 in FIG. 1.

[0111] Prior to the cyanidation 444, the solids from the separation 438can be subjected to an optional slaked lime boil step 443 to facilitatethe recovery of silver during the cyanidation 444 by the decompositionof silver jarosite compounds formed during the total oxidative leach414.

[0112] The precious metals are in the solids remaining after theseparation 438. Now that pyrite and other encapsulating minerals in theoriginal concentrate have been decomposed, the precious metals areamenable to cyanidation 444.

[0113] In the cyanidation step 444, the solids are leached with NaCNunder alkaline conditions. In order to effect this, the solids areslurried up with cyanide solution to form a 30-40% solids slurry.Additional NaCN and slaked lime are added as required to maintain aminimum NaCN concentration of about 0.2 to about 0.5 g/l NaCN, with a pHof about 10. The temperature is ambient and usually about 4 to 8 hoursretention time is required in continuous mode of operation.

[0114] Both gold and silver report in high yield to the cyanidesolution, and are recovered typically by the established process ofcarbon-in-pulp circuit, whereby activated carbon is added to the cyanideslurry to absorb the precious metals, without the necessity offiltration. The loaded carbon, now rich in precious metals is separatedby screening (445) and the barren pulp discarded to tailing.

[0115] The loaded carbon is treated by established methods to recoverthe precious metals content by a leach/electrowin/smelt process (447).The product is generally Dore metal containing both gold and silver,which is sent to a gold refinery 449 for final separation of gold fromsilver. Barren carbon from a carbon regeneration step 451 after theprecious metals recovery, is recycled to the carbon-in-pulp circuit 444.

[0116] The overall recovery of precious metals by the total process isgenerally well over 90%, and under optimum conditions approach 99%.

[0117] A test was carried out in which desulphurized residue wasprocessed in a total oxidative leach 414 at 220° C. for 2 hours underoxygen pressure and then depressurized and cooled to room temperature.The resultant slurry was neutralized to pH 3 with limestone and thenfiltered. The filtered cake was then leached with cyanide solution understandard conditions to leach gold and silver.

[0118] The gold extraction after the total oxidative leach 414 andcyanidation 444 was 97% with only 1.0 kg/t NaCN consumption. Incomparison, the gold extraction on a residue that had not been oxidizedin the total oxidative leach 414 was only 34% and cyanide consumptionwas extremely high at 19.0 kg NaCN/t.

[0119]FIG. 4 is a flow diagram of Mode A. Steps which correspond withthose of the embodiment of FIG. 1 are given the same reference numerals.

[0120] The process comprises a pressure oxidation stage 12 in whichsulphide minerals in the concentrate or ore are oxidized by highpressure oxygen, followed by a liquid/solid separation (e.g. filtration)24, producing a solid (pressure oxidation filter cake) 25 and pressureoxidation filtrate 29.

[0121] The solid 25 contains all or almost of the copper content of thefeed concentrate, and is treated for copper recovery 14 by acidleaching, solvent extraction and electrowinning as in the embodiment ofFIG. 1, thus producing high quality copper cathodes, and a residue 35which may contain precious metals. The residue 35 can be treated forprecious metal recovery, as described with reference to FIGS. 3A and Babove. This is indicated by the block 155 in FIG. 4.

[0122] The filtrate 29 is purified at 500 to remove deleterious elementssuch as Cu, Fe and Zn, by neutralization with slaked lime to about pH 6,as described with reference to FIG. 1, producing a purified solution 36,after filtration, containing Ni, Co and certain other elements such asMg which may be present in the feed concentrate.

[0123] The solution 36 is treated for Ni/Co recovery as described withreference to FIG. 1. This is indicated by the block 38 in FIG. 4. Thesolution 39 produced in 38 is recycled back to the pressure oxidation12, to complete the cycle, as before (stream 516 in FIG. 1).

[0124] While only preferred embodiments of the invention have beendescribed herein in detail, the invention is not limited thereby andmodifications can be made within the scope of the attached claims.

What is claimed is:
 1. A process for the recovery of nickel or cobaltvalues from a concentrate containing nickel or cobalt hydroxide,comprising the steps of: subjecting the concentrate to a leaching stagewith an ammonium solution to produce a leach solution containing nickelor cobalt values and a residue; and controlling the concentration ofnickel in the leach solution to a maximum value of about 3 to 25 g/l. 2.The process according to claim 1, wherein the maximum value is fromabout 8 to 15 g/l.
 3. The process according to claim 2, wherein themaximum value is about 10 g/l.
 4. The process according to claim 1,wherein the leaching stage is effected with an ammonium sulphatesolution.
 5. The process according to claim 4, wherein the leachingstage is effected at a pH of from about 6 to
 8. 6. The process accordingto claim 4, wherein the ammonium sulphate solution has a concentrationof from about 150 to 250 g/l.
 7. The process according to claim 6,wherein the ammonium sulphate solution has a concentration of about 200g/l.
 8. The process according to claim 1, wherein the nickel or cobaltleaching stage is effected with an ammonium carbonate solution.
 9. Theprocess according to claim 1, wherein the nickel or cobalt leachingstage is effected with a mixture of ammonium sulphate and ammoniumcarbonate in solution.
 10. The process according to claim 1, furthercomprising the steps of: subjecting the residue to an acidic washingstage to produce a wash solution containing nickel or cobalt values anda discardable residue; subjecting the wash solution to a selectiveprecipitation treatment to obtain a solid containing nickel or cobalthydroxide; and recycling the solid to the leaching stage.
 11. Theprocess according to claim 1, further comprising the steps of:subjecting the residue to an acidic washing stage to produce a washsolution containing nickel or cobalt values and a discardable residue;and treating the wash solution for the recovery of the nickel or cobaltvalues therefrom.
 12. The process according to claim 10, furthercomprising the step of subjecting the residue to a washing stage priorto the acidic washing stage to produce a second wash solution containingnickel or cobalt values and a residue which is subjected to the acidicwashing stage.
 13. The process according to claim 12, further comprisingthe step of subjecting one or both of the leach solution and the furtherwash solution containing nickel or cobalt values to solvent extractionto recover nickel or cobalt values therefrom.
 14. The process accordingto claim 13, wherein the solvent extraction is effected with a nickelextractant to produce a nickel containing solution.
 15. The processaccording to claim 13, wherein the solvent extraction is effected with acobalt extractant to produce a cobalt containing solution.
 16. Theprocess according to claim 12, wherein the solvent extraction comprisesthe steps of: effecting a cobalt solvent extraction in the presence ofmagnesium ions with a cobalt extractant to produce a cobalt extractantloaded with cobalt ions and a first raffinate containing nickel andmagnesium ions in solution; effecting a magnesium solvent extraction onthe first raffinate with a magnesium extractant to produce a magnesiumextractant loaded with magnesium and cobalt ions and a second raffinate;and effecting a nickel solvent extraction on the second raffinate with anickel extractant to produce a nickel loaded extractant and a thirdraffinate.
 17. The process according to claim 16, further comprising thestep of stripping the cobalt and nickel loaded extractants to producecobalt and nickel solutions, respectively.
 18. The process according toclaim 17, further comprising the step of subjecting the cobalt andnickel solutions to electrowinning to recover cobalt and nickeltherefrom.
 19. The process according to claim 17, further comprising thestep of: stripping the magnesium extractant to produce a pregnantsolution containing magnesium and cobalt ions; and recycling thepregnant solution to the selective precipitation treatment.
 20. Theprocess according to claim 16, wherein the cobalt extractant is the sameas the magnesium extractant, the extractant being more selective forcobalt than for magnesium.
 21. The process according to claim 13,wherein the solvent extraction comprises the steps of: effecting acobalt solvent extraction at a pH of about 6 to 8 to produce a cobaltsolution and a first raffinate; and effecting a nickel solventextraction on the first raffinate at substantially the same pH as thecobalt solvent extraction to produce a nickel solution and a secondraffinate.